Production of biodiesel and glycerin from high free fatty acid feedstocks

ABSTRACT

A system and method for the conversion of free fatty acids to glycerides and the subsequent conversion of glycerides to glycerin and biodiesel includes the transesterification of a glyceride stream with an alcohol. The fatty acid alkyl esters are separated from the glycerin to produce a first liquid phase containing a fatty acid alkyl ester rich (concentrated) stream and a second liquid phase containing a glycerin rich (concentrated) stream. The fatty acid alkyl ester rich stream is then subjected to distillation, preferably reactive distillation, wherein the stream undergoes both physical separation and chemical reaction. The fatty acid alkyl ester rich stream is then purified to produce a purified biodiesel product and a glyceride rich residue stream. The glycerin rich second liquid phase stream may further be purified to produce a purified glycerin product and a (second) wet alcohol stream. Neutralization of the alkaline stream, formed during the alkali-catalyzed transesterification process, may proceed by the addition of a mineral or an organic acid.

This application is a continuation-in-part of U.S. application Ser. No.13/307,871 filed Nov. 30, 2011, which is a continuation of U.S.application Ser. No. 11/504,828, filed Aug. 15, 2006, which is acontinuation-in-part application of U.S. application Ser. No.10/766,740, filed on Jan. 26, 2004, which claims the benefit of U.S.Patent Application Ser. Nos. 60/443,049, filed Jan. 27, 2003, and60/537,251, filed Jan. 15, 2004, each of which is hereby incorporated.

FIELD OF THE INVENTION

The present invention relates to improved processes and systems forbiodiesel production.

BACKGROUND OF THE INVENTION

There is continued and growing interest in the use of renewableresources as replacements for petroleum-derived chemicals. Fatty acidalkyl esters (FAAEs) produced from fats and oils have been investigatedas replacements for such petroleum-derived materials, particularlydiesel fuel.

It has long been known that triglycerides from fats and oils can be usedas fuels for diesel engines. However, such use typically results inengine failure. Remedies for such engine failure wherein conversion offatty acids, found in lipids, into simple esters, such as methyl andethyl esters, has been proposed. See, for instance, the processdescribed in U.S. Pat. No. 6,398,707. An increasing body of evidenceindicates that these esters perform well in essentially unmodifieddiesel engines and that such esters may effectively reduce the output ofparticulate and hydrocarbon pollutants relative to petroleum-dieselfuel. The term “biodiesel” is applied to these esters.

Processes for biodiesel production have been known for many years. Forinstance U.S. Pat. No. 4,164,506 discloses a biodiesel synthesis whereinfatty acids are subjected to acid catalysis. The conversion oftriglycerides with base catalysis is described in U.S. Pat. Nos.2,383,601 and 2,494,366. Conversion of both free fatty acids andtriglycerides with enzyme catalysis is disclosed in U.S. Pat. Nos.4,956,286, 5,697,986 and 5,713,965. None of these processes, however,completely addresses the production of biodiesel from low value highfree fatty acid feedstocks.

An economic analysis of any process for the production of biodieselindicates that feedstock cost is the largest portion of production costfor biodiesel. Whereas a 15 weight percent free fatty acid (FFA)feedstock is the highest content that any contemporary commercialprocess has proposed to handle, producers (in order to conserve costs)would prefer to use feedstocks having up to 100 weight percent FFAcontent.

Further, most of the processes of the prior art are unattractive becausethey rely upon acid catalyzed esterification of fatty acids. Acidcatalysis is not suitable for processing such feedstocks containing FFAconcentrations for two principal reasons. First, an excessive amount ofacid catalyst is required in order to fully convert feedstocks havinghigh FFA content. Since the acid catalyst must be neutralized beforeprocessing the glycerides, the increased catalyst loading results in anexcessive amount of generated salt. Further, such processes generate alarge volume of waste water as disclosed in U.S. Pat. Nos. 4,303,590,5,399,731 and 6,399,800.

While enzymatic catalysis has been reported in the literature foresterification of free fatty acids, it is disadvantageous because ofreaction product inhibition from the presence of water which resultswhen the free fatty acids in the feedstock are esterified with enzymes.Another problem evidenced from enzymatic processing is the high cost ofenzymatic catalysts. Further, enzymatic catalysts have a limited life.

To avoid two-phase operation in packed bed and other reaction settings,some conventional processes for biodiesel production use volatile, toxicco-solvents. Such a process is disclosed in U.S. Pat. No. 6,642,399 B2.The use of volatile, toxic co-solvents is environmentally unacceptable.

Further, some prior art processes for producing biodiesel employ waterto wash residual glycerin and salts from the FAAEs. This, however,generates a large volume of wastewater and increases the risk of formingFAAE emulsions, as disclosed in U.S. Pat. No. 5,399,731.

To be economically profitable, the biodiesel industry must takeadvantage of lower cost feedstocks. Yield is a very important criterionas feedstock costs approach two thirds of the total cost of productionof biodiesel. To gain market share in the fuels industry, biodiesel mustbe competitively priced with conventional hydrocarbon diesel.

Alternative processes need to be developed which do not require highpressures or acid catalysis. In addition, such processes should notemploy toxic co-solvents or water for the extraction of impurities. Suchprocesses also need to produce high yield of biodiesel as well as employinexpensive feedstocks. Further, such feedstocks need to have a high FFAcontent in order to be competitive with petrodiesel.

SUMMARY OF THE INVENTION

A process is disclosed which combines several unit operations into aneconomical and unique process for the conversion of free fatty acids toglycerides and the subsequent conversion of glycerides to glycerin andFAAEs. The fatty acid alkyl esters of the invention produced inaccordance with the invention are typically fatty acid methyl estersthough other fatty acid alkyl esters may be produced.

The invention relates to a process for converting low-value, high freefatty acid (FFA) feedstocks to biodiesel and high quality glycerin at amarket price comparable to that of petroleum derived diesel fuels. Theprocess of the invention therefore substantially departs fromconventional concepts and designs of the background art. In so doing,the inventive process provides a process and apparatus primarilydeveloped for the purpose of producing fatty acid alkyl esters and highquality glycerin from any low-value high free fatty acid feedstock.

Another aspect of the invention relates to separation and purificationof major by-products of biodiesel production to render glycerin at apurity level greater than 95 or 99.7 percent, with non-detectable levelsof alcohol and less than 0.5 percent weight/weight (w/w) salts.

The invention further relates to minimization of waste streams duringnormal operations, the use of lower operating temperatures and pressuresthan other commercial biodiesel processes, the non-use of toxicco-solvents and the production of a high quality glycerin byproduct.

In a preferred embodiment, the process is a continuous process.

The major steps of the process include the transesterification of aglyceride stream with an alcohol, preferably in the presence of basecatalyst, to convert the glycerides to fatty acid alkyl esters andglycerin.

The fatty acid alkyl esters are then separated from the glycerin toproduce a first liquid phase containing a fatty acid alkyl ester richstream and a second liquid phase containing a glycerin rich stream.

The fatty acid alkyl ester rich stream is then subjected todistillation, preferably reactive distillation, wherein the streamundergoes both separation and chemical reaction. By means of reactivedistillation, the stream is separated into a bottoms fraction containinga plurality of the fatty acid alkyl esters and an overhead fraction(principally of alcohol, a first wet alcohol stream), whilesimultaneously chemically reacting two or more stream componentstogether in such a way as to remove unwanted impurities in one or moreoutput stream(s). Such reactive distillation for example increases theyield amount of glycerides exiting the distillation column whileincreasing the purity of the biodiesel exiting the distillation column.The combination of chemical reaction and separation offers distinctadvantages over conventional processes.

The fatty acid alkyl ester rich stream is then purified to produce abiodiesel product and a glyceride rich residue stream.

The glycerin rich second liquid phase stream may further be purified toproduce a purified glycerin product and a (second) wet alcohol stream. Aportion of the purified glycerin product may then be recycled into aglycerolysis reactor (in a glycerolysis process described in more detailbelow) for reaction with the free fatty acids.

The wet alcohol streams may further be purified, preferablycontinuously, to produce a purified alcohol product. Further, at least aportion of the purified alcohol product may be recycled into thetransesterification reactor for reaction with the glycerides.

Neutralization of the alkaline stream, formed during thealkali-catalyzed transesterification process, may proceed by theaddition of a mineral acid or more preferably an organic acid to thestream. Neutralization may occur by addition of the acid to thetransesterification effluent stream directly or to the fatty acid alkylester rich stream and/or glycerin rich stream after such streams havebeen separated from the transesterification effluent stream.

BRIEF DESCRIPTION OF THE DRAWINGS

The features of the invention will be better understood by reference tothe accompanying drawings which illustrate presently preferredembodiments of the invention. In the drawings:

FIG. 1 is a schematic flow diagram of the process of the invention.

FIG. 2 is a schematic block diagram of the biodiesel production systemin accordance with the invention;

FIG. 3 is a schematic block diagram showing the basic steps of theproduction of biodiesel in accordance with the process of the invention;

FIG. 4 is a schematic flow diagram of the process of the inventionwherein a mineral acid is used in the neutralization of the alkalicatalyst used during transesterification;

FIG. 5 is a schematic flow diagram of the process of the inventionwherein an organic acid is used in the neutralization of the alkalicatalyst used during transesterification;

FIG. 6 is a schematic block diagram which demonstrates reactivedistillation of a fatty acid alkyl ester rich stream upon separationfrom the transesterification effluent stream, as set forth in ExampleNo. 6; and

FIG. 7 is a schematic flow diagram showing an alternate embodiment ofthe process having a modified glycerolysis reactor vent system.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

In the process of the invention, biodiesel is prepared by reactingglycerides with an alcohol in a transesterification reactor to producefatty acid alkyl esters. This reaction typically occurs in the presenceof an alkali catalyst. The alcohol is typically a C₁-C₅ alcohol,preferably methanol.

The resulting transesterification effluent stream may then be separatedinto a fatty acid alkyl ester rich stream and a glycerin rich stream.Each of these streams may then be purified to maximize the efficiency inrecovery of biodiesel, glycerin and alcohol.

The alkaline transesterification effluent stream formed during thealkali-catalyzed transesterification process may be directly treatedwith a neutralizing agent, such as a mineral acid or an organic acid.Alternatively, the neutralizing agent may be added to the fatty acidalkyl ester rich stream and/or the glycerin rich stream after thestreams have been separated from the transesterification effluentstream. Fatty acid alkyl esters are recovered from this pH adjustedstream.

Subsequent to neutralization, the neutralized stream may further bepurified, such as by distillation or fractionation.

The process of the invention may further consist of an esterificationstep wherein a free fatty acid feedstock is first converted toglycerides. The resulting glycerides are then introduced into thetransesterification reactor.

The use of the acid as neutralizing agent converts soaps, formed in thetransesterification reactor, to free fatty acids. The soap forms fromthe action of caustic with fatty acids in the transesterificationreactor. The presence of the soap makes it very difficult to effectuatephase separation between the fatty acid alkyl esters and the solution ofglycerin, water, alcohol and salt. As a result, the soap emulsifies andretains much of the fatty acid alkyl esters in the glycerin rich phase.Purification of the glycerin rich phase is therefore complicated by thepresence of the soap and the yield of alkyl esters is decreased.

An overview of the process of the invention may be presented in FIG. 3wherein a feedstock 1 containing free fatty acids is introduced into aglycerolysis reactor 2 with glycerin wherein the free fatty acids areconverted to glycerides. The glycerides are then introduced intotransesterification reactor 4 with alcohol wherein the glycerides aretransesterified to form fatty acid alkyl esters and glycerin. Thealcohol/alkali stream 3 may be introduced into transesterificationreactor 4 as a combined mixture of alkali catalyst and alcohol, oralternatively the alkali catalyst and alcohol may be introduced into thetransesterification reactor as separate streams into transesterificationreactor 4. The transesterification effluent stream or a portion thereofis then neutralized during neutralization/phase separation step 5,either before or after the effluent stream is separated into a fattyacid alkyl ester rich stream and a glycerin rich stream. Ultimately,alcohol, glycerin and biodiesel may be refined in alcohol refining step6, glycerin refining step 7 and biodiesel refining step 8, respectively.The alcohol typically exits the system as a small portion of wastestream 9 or is recycled via flow 11 back to the transesterificationreactor. Refined glycerin is isolated in technical grade glycerin stream13 or may be recycled back via flow 15 to the glycerolysis reactor. Thealkyl esters may further be refined in biodiesel refining step 8 toproduce purified biodiesel stream 18 or may exit the system as a portionof waste stream 19 which may be useful, for example, as a burner fuel.

The process of the invention may be a continuous process wherein one ormore of the following steps are carried out in a continuous fashion:

(1) the optional conditioning of a fatty acid containing feedstock byheating, mixing and filtering;

(2) continuously reacting the free fatty acids in the feedstock withglycerin in a glycerolysis reactor to produce glycerides;

(3) reacting the glycerides in a transesterification reactor withalcohol to render fatty acid alkyl esters and glycerin. This reactionpreferably occurs in the presence of an alkali catalyst;

(4) separating fatty acid alkyl esters and glycerin from thetransesterification effluent stream to yield a fatty acid alkyl esterrich stream and a glycerin rich stream, optionally separating the fattyacid alkyl ester and glycerin from the transesterification effluentstream by reactive distillation wherein a reaction in the distillationor fractionation column assists in the separation of the fatty acidalkyl esters and glycerin;

(5) purifying the fatty acid alkyl ester rich stream and recoveringalcohol from the stream. The resultant purified fatty acid alkyl esteris acceptable for use as biodiesel;

(6) purifying the glycerin rich stream, preferably by use of an organicacid, such as a weak organic acid like acetic acid, formic acid orpropionic acid, and recovering alcohol from the stream. The purifiedglycerin may then be introduced into the glycerolysis reactor;

(7) purifying the wet alcohol streams resulting from steps (5) and (6)above and removing water from the streams; and

(8) recycling at least a portion of the purified alcohol to thetransesterification reactor for reaction with the glyceride.

The feedstock, from which the biodiesel may be produced, typicallycontains a plurality of free fatty acids. The feedstock typicallycontains between from about 3 to about 100 weight percent of free fattyacids and, optionally, a fat and/or an oil.

Typically, the feedstock is a lipid feedstock. The free fatty acidfeedstock for use in the invention may be a low-grade lipid materialderived from animal fats and vegetable oils, including recycled fats andoils. For instance, the feedstock for the production of biodiesel fuelmay be a grease feedstock, such as a waste grease or a yellow grease.Such low-grade lipid materials are very complex and typically aredifficult to economically process using current state of the artprocesses because of their high free fatty acid levels (ranging from afew percent to 50 percent, and higher). In addition, such materialscontain unprocessable material and contaminants that must be removedprior to processing or during refinement of the products.

The feedstock may be first introduced into a conditioning vessel orreactor that is operative to heat, mix and/or filter the feedstock toproduce a conditioned feedstock. The feedstock may then be filtered,such as by using a traveling screen.

Subsequent to filtration, the concentration of free fatty acids in theconditioned lipid feedstock may be measured. Optionally, theconcentration of free fatty acids in the conditioned feedstock may becontinuously measured throughout the process. Measurements may be madewith an in-line free fatty acid titration device that is operative toquantify the concentration of the free fatty acid in the conditionedfeedstock.

During conditioning, the feedstock may be heated to a temperature in therange of about 35° C. to about 65° C., preferably between from about 55°C. to about 65° C., while mixed. A uniform mixture of glycerides, freefatty acids and unsaponifiable materials are typically present in theconditioned feedstock.

During glycerolysis, glycerin is used as a reactant to convert the freefatty acids in the feedstock to glycerides (mono-, di-, andtriglyceride). Reaction of the free fatty acids in the feedstocktypically occurs in the absence of a catalyst. In the glycerolysisreactor, the free fatty acid in the feedstock is mixed and continuouslyreacted with glycerin at an appropriate temperature and pressure torender a glycerolysis reactor effluent stream that contains generallyless than about 0.5 percent by weight of free fatty acids and aplurality of glycerides. Glycerolysis typically occurs in the absence ofboth catalyst and co-solvent.

The glycerin, typically a purified glycerin product, is normally addedto the glycerolysis reactor at a rate that is greater than thestoichiometric amount of glycerin required for the glycerolysisreaction. The amount of glycerin introduced to the glycerolysis reactoris generally in a stoichiometric proportion of about 3:1 of free fattyacid to glycerin in order to render the glyceride. In a preferredembodiment, the amount of glycerin added to the glycerolysis reactor isat a rate in the range of about 35 percent to about 400 percent of thestoichiometric amount of free fatty acids in the feedstock.

Preferably, glycerolysis is conducted at a temperature in the range ofabout 150° C. to about 250° C., typically from about 180° C. to about250° C., more typically from about 180° C. to 230° C. The reactiontypically proceeds under agitation. The reaction is further typicallyconducted at a pressure of about 0.1 pounds per square inch absolute toabout 7 pounds per square inch absolute.

Reaction of the free fatty acids and glycerin typically occurs in theabsence of a catalyst. The glycerolysis reactor effluent stream maycontain less than 0.5 percent by weight of free fatty acids and aplurality of glycerides.

The glycerolysis is typically a continuous reaction. The continuousreaction of the free fatty acid in the feedstock with glycerin toproduce the glyceride in the glycerolysis reactor may be conducted inresponse to a signal from the in-line fatty acid titration device.

During glycerolysis, water is removed; the produced glycerides beingessentially water-free. Water is typically continuously removed from theglycerolysis reactor 34 as a vapor through a fractionation column or avent in the reactor headspace. Preferably, the vapor vented from theglycerolysis reactor 34 is fractionated to yield a liquid fractionhaving a high concentration of glycerin and a vapor fraction or secondliquid fraction having a high concentration of water. The liquidfraction containing the glycerin may then be returned to theglycerolysis reactor 34.

Another embodiment of the glycerolysis reactor vent system is presentedin FIG. 7, process 200 wherein the feedstock first undergoes a heatingstep (corresponding to FIG. 1 step 32), followed by a glycerolysis step(corresponding to FIG. 1 step 34) in which free fatty acids areconverted to glycerides, and then a glycerolysis effluent cooling step(corresponding to FIG. 1 step 38). One purpose of the glycerolysisreactor vent system is to remove water from the reactor as it isproduced by the glycerolysis reactions. Removing said water allows thereactions to proceed beyond the equilibrium limitations that wouldotherwise constrain conversion. However, glycerin and/or feedstock canbe evaporated and/or entrained (e.g., as droplets or an aerosol) intothe vapor stream that is vented from glycerolysis reactor unit 34.

A significant portion of the glycerin and/or feedstock in the vent vaporstream can be recovered by condenser 205 and returned to theglycerolysis step 34, as explained above. In this embodiment, condenser205 is preferably operated such that very little water vapor iscondensed and returned to the glycerolysis step 34 in order to minimizethe water content in the glycerolysis reactor 34. As a consequence, asignificant amount of glycerin and/or feedstock can pass throughcondenser 205. In some embodiments, the remaining glycerin and/orfeedstock that pass through 205 may be recovered by a second condenser210, which is also operated such that the majority of the water vaporremains in vapor state. Again, this helps to minimize the amount ofwater that is recycled back into the glycerolysis process. In order toenhance the recovery of feedstock and/or glycerin from the vapor stream,the condensers 205, 210 may be operated in such a way so as to collectand/or coalesce liquid droplets and aerosols in addition to collectingcondensable vapor components.

Either or both of the liquid streams from condenser units 205 and 210may be either collected in phase separation unit 215 or returneddirectly to the glycerolysis reactor 34. In another embodiment, a singlecondenser may be used instead of condensers 205 and 210 to recoverfeedstock and glycerin that leave the glycerolysis reactor 34 in thevapor stream. Again, the condensate may be either collected in phaseseparation unit 215 or returned directly to the glycerolysis reactor 34.In one embodiment, the condensers 205, 210 may operate at conditionssuch that more or less water is condensed and returned back to theglycerolysis step 34, although it is generally preferable to minimizethe amount of water that is reintroduced to the glycerolysis reactor.

Preferably, the liquid collected in phase separation unit 215 separatesinto a heavy phase having a high concentration of glycerin with somewater and a light phase having a high concentration of feedstock. Thelight phase containing the feedstock may then be returned to theglycerolysis reactor unit 34, thus reducing feedstock waste andmaximizing fatty acid alkyl ester yield. The heavy phase containingglycerin and some water is sent to glycerin processing 220 with thecrude glycerin produced in the transesterification process.

A water-rich vapor stream 225 is ultimately condensed in unit 230 toyield a water-rich liquid stream 235. A component-rich stream (e.g.water-rich stream, glycerin-rich stream, feedstock-rich stream, etc. . .. ) is a stream having an higher concentration as compared to theearlier stream, and may be referred to as a rich stream, an enrichedstream, or a stream having a high concentration of a particularcomponent. The water-rich vapor stream 225 and condensed water-richliquid stream 235 may contain alcohol released from the glycerolysisreactor unit 34 (alcohol may be produced during glycerolysis if thefeedstock contains any alkyl esters). A final condenser unit 240 cooledwith chilled fluid may also be included to remove any remainingcondensable components in the vapor stream to protect the vacuum unit245. Any condensate from optional unit 240 may be combined with thewater-rich liquid stream 235 from unit 230. Any liquids, aerosols, orrecoverable vapors remaining in the vapor stream after unit 240 aretrapped in a vent scrubber between the vacuum unit outlet and theatmosphere.

The glycerolysis reactor 34 may consist of two or more continuousstirred tank reactors operated in series. The residence time of suchreactors is typically not more than about 500 minutes, and preferablynot more than 200 minutes.

A plurality of glycerides contained in the glycerolysis effluent streamare reacted with an alcohol in the transesterification reactor, such asa continuous stirred tank reactor. In this reaction, the glycerides inthe glycerolysis reactor effluent stream are transesterified into fattyacid alkyl esters and glycerin. Transesterification proceeds at anappropriate temperature and pressure to produce the desiredtransesterification reactor effluent stream.

Transesterification, which preferably is a continuous process, occurs inthe presence of a base catalyst. Suitable base catalysts include suchalkali catalysts as potassium hydroxide and sodium hydroxide. The alkalicatalyst may be added to the transesterification reactor at a ratesufficient to catalyze the reaction. Typically, the amount of alcoholadded to the transesterification reactor is from about 0.5 percent byweight to 2.0 percent by weight of the glycerides present in theglycerolysis effluent stream.

Alternatively, an alkoxide, such as potassium methylate, may be added tothe transesterification reactor to facilitate the base catalysis. Assuch, the rapid conversion of glycerides to alkyl esters may occur inthe presence of caustic alkoxide, such as caustic methoxide catalysts.

The transesterification reaction typically occurs at a temperature inthe range of about 25° C. to about 65° C., preferably from about 55° C.to about 65° C., and at a pressure of about 14.5 psia to about 3,625psia.

The alcohol is normally added to the transesterification reactor at arate that is greater than the stoichiometric amount of alcohol requiredfor the alkali catalyzed transesterification reaction. For instance, thealcohol may be added to the transesterification reactor at a rate equalto about 200 percent of the stoichiometric amount of alcohol requiredfor the catalyzed reaction.

Preferably, multiple alcohol or catalyst additions are made to thetransesterification reactor.

The transesterification reactor typically contains at least twocontinuous stirred tank reactors that are operated in series. Each ofthe tank reactors typically has a residence time of about 15 to about 90minutes, typically about 60 minutes.

The resulting transesterification reactor effluent stream contains afatty acid alkyl ester and glycerin. Preferably, at least a portion ofthe glycerin is removed from the transesterification reactor before theplurality of glycerides are reacted with the alcohol.

A plurality of the resulting fatty acid alkyl esters may then beseparated from the glycerin in the transesterification effluent stream.Separation into two distinct immiscible phases, i.e., a first liquidphase in which the plurality of fatty acid alkyl esters may beconcentrated and a second liquid phase in which glycerin may beconcentrated, is typically dependent upon the differences in densitiesin the two phases and employs gravitational force and/or centrifugalforce.

Typically, the two phases are separated at a temperature of about 25° C.to about 65° C. to produce the fatty acid alkyl ester rich stream andglycerin rich stream. This separation process may be a continuousoperation and may be performed in a clarifier or by means of membranefiltration.

In a preferred embodiment, the fatty acid alkyl ester rich stream issubjected to reactive distillation to separate the fatty acid alkylester rich stream into a bottoms fraction, an overhead fraction(principally comprising excess alcohol) and a fatty acid alkyl esterproduct stream. Such separation utilizes the differences in the vaporpressures of the components of the fatty acid alkyl ester rich streamand the reactive loss of glycerin. The conditions in the distillation orfractionation column including temperature and pressure conditions,simultaneously with and in the same vessel wherein the said separationoccurs, promote a chemical reaction to occur. Reactive distillation inthe embodiment depicted in FIG. 6 decreases the concentration ofglycerin and increases the amount of glycerides exiting the column.Thus, reactive distillation increases the efficiency of the productionprocess.

The end result of reactive distillation is that the amount of glycerinseen in the transesterification effluent stream, or the first liquidphase, is greater than the total amount of glycerin which exits thedistillation or fractionation column. This is attributable to thereaction of the glycerin with free fatty acids and or fatty acid alkylesters in the reactive distillation column to form glycerides.

Preferably, the overhead fraction produced by the fatty acid alkyl esterdistillation column is a (first) alcohol stream which comprisesessentially the alcohol. Preferably the bottoms fraction comprisesimpurities having a high boiling point, unsaponifiable materials,monoglycerides, diglycerides, triglycerides and fatty acids. Preferably,the fatty acid alkyl ester product produced by the fatty acid alkylester distillation column meets ASTM specification D 6751.

Preferably, the fatty acid alkyl ester distillation column orfractionation column is operated at a pressure below about 2 pounds persquare inch absolute. More preferably, the fatty acid alkyl esterdistillation column or fractionation column is operated at a pressure inthe range of about 0.1 pounds per square inch absolute to about 2 poundsper square inch absolute. Preferably, the fatty acid alkyl esterdistillation column or fractionation column is operated at a temperaturein the range of about 180° C. to about 280° C., more preferably betweenfrom about 180° C. to about 230° C. Preferably, the fatty acid alkylester distillation column or fractionation column contains a packingmaterial.

The glycerin rich second liquid phase stream may further be purified andalcohol recovered from it. The recovered alcohol is operative to producea purified glycerin product and a (second) wet alcohol stream. In apreferred embodiment, this step employs one or more of glycerinfractionation (wherein the fractions within the glycerin rich stream areseparated by distillation), phase separation (wherein the impuritiesthat co-fractionate with glycerin are removed by immiscibility anddifferences in density) and glycerin polishing (wherein other impuritiesare removed from glycerin).

The glycerin rich stream may further be subjected to phase separationwherein a fatty acid alkyl ester rich liquid phase and a glycerin richliquid phase are separated and the two liquid phases may then be subjectto purification as described in the paragraphs above.

The glycerin rich stream may further be purified in a glycerindistillation or fractionation column to produce a bottoms material, aside stream and an overhead stream. Preferably, the bottoms materialcontains essentially waste materials; the side stream containsessentially glycerin and trace impurities; and the overhead streamcontains essentially alcohol and water that is collected for furtherpurification and recycled.

Preferably, the glycerin distillation column is operated at an elevatedtemperature between about 180° C. and about 280° C., more preferablybetween from about 180° C. to about 230° C. The distillation column istypically operated at a reduced pressure, of below about 2 pounds persquare inch absolute, typically the pressure is in the range of about0.1 pounds per square inch absolute to about 2 pounds per square inchabsolute.

The glycerin rich stream may further be subjected to a decolorizationcolumn wherein colored impurities and odors are removed from theglycerin, i.e., “glycerin polishing”. The decolorization columntypically comprises a packed bed of activated carbon operated at atemperature in the range of about 35° C. to about 200° C., preferablybetween from about 40° C. to about 100° C. The contact time is generallyless than four hours. Activated carbon fines carried through the packedbed are removed by filtration.

Water may further be removed from the wet alcohol streams to renderpurified alcohol by subjecting the wet alcohol stream to an alcoholdistillation or fractionation column at a temperature in the range ofabout 60° C. to about 110° C. and at a pressure in the range of about 14pounds per square inch absolute to about 20 pounds per square inchabsolute. Preferably, this purification comprises adsorption ontomolecular sieves that can then be dried and reused or distillationresulting in a bottoms product consisting mainly of water.

At least a portion of the purified glycerin product may then be returnedto the glycerolysis reactor for reaction with free fatty acids in thefeedstock; at least a portion of the purified alcohol being recycledinto the transesterification reactor for reaction with glycerides.

It is typically desired to neutralize the fatty acid alkyl ester andglycerin produced in the transesterification reactor. Neutralization isoften required in light of the caustic conditions which characterizetransesterification. Such neutralization may occur by addition of anacid to the transesterification effluent stream or to either the fattyacid alkyl ester rich stream or glycerin rich stream after such streamsare separated from the transesterification effluent stream. Suitableacid treatments include mineral or more preferably organic acidtreatments.

Suitable mineral acids include sulfuric acid and phosphoric acid.Reaction of the alkali catalyst with a mineral acid renders an insolublesalt that is removed from the glycerin rich stream in a solidsseparation operation. FIG. 4 is illustrative of the process wherein amineral acid, such as phosphoric acid, is employed. In particular, FIG.4 illustrates introduction of feedstock 310 containing free fatty acidsinto glycerolysis reactor 312 wherein the free fatty acids are convertedto glycerides by esterification. The glycerides are then introduced intotransesterification reactor 314 with alcohol 316 and alkali catalyst 318wherein the glycerides are transesterified to form fatty acid alkylesters and glycerin.

The transesterification effluent stream is first separated in 1st phaseseparation 320 into a fatty acid alkyl ester rich stream and a glycerinrich stream. Each of these streams may then be purified in 2nd phaseseparation 322 in accordance with the processes described herein.

The neutralization acid, phosphoric acid, 324 is added either prior to1st phase separation 320 or subsequent to 1st phase separation 320 ofthe transesterification effluent stream after the fatty acid alkyl esterrich stream and glycerin rich stream have been separated. Suchalternative or combination ports of introduction of the acid into theprocess are represented by the dotted lines in FIG. 4.

Unfortunately, use of phosphoric acid renders an insoluble precipitate.The formation of the insoluble precipitate mandates the use of filter326 and/or filter 328. Suitable filters include rotary vacuum drumfilters, plate and frame presses as well as belt presses.

In addition to the use of a filtration unit, use of a mineral acidfurther requires the rinsing of the insoluble by-product salts in orderto wash residual organic materials from them. Suitable solvents includeC₁-C₅ alcohols, such as methanol. Illustrated in FIG. 4 is theintroduction of alcohol solvent 329 for use as alcohol rinse 330 whichremoves organic residue from the filter cake. Vacuum dry 332 is thenused to remove alcohol from the filter cake and to dry the purified saltwhich then exits the process as waste 334. The solvent may then berecovered as stream 364 for reuse in the process.

Preferably, the process comprises drying the insoluble salt in a dryerunder conditions wherein the temperature of the dryer exceeds theboiling point of the solvent at the operating pressure of the dryer. Thedryer may optionally be operated under a vacuum to improve the drying.The dryer may further include a condenser to recover the solvent forreuse.

FIG. 4 further illustrates the refining of alcohol, glycerin andbiodiesel in alcohol refinery vessel 346, glycerin refinery vessel 348and biodiesel refinery vessel 350, respectively. The alcohol typicallyexits the system as byproduct stream 352 or is recycled via 360 back tothe transesterification reactor. Refined glycerin is isolated aspurified glycerin 354 or may be recycled back at 362 to the glycerolysisreactor. The alkyl esters may further be purified to produce purifiedbiodiesel 356 or may exit the system as byproduct 358 in the form of,for example, burner fuel.

It is more preferable to employ an organic acid versus a mineral acid,however. While there are inorganic acids that don't create precipitatingsalts upon neutralization with the transesterified stream, all sufferfrom serious disadvantages. For instance, hydrochloric and perchloricacid produce chlorides in the process streams which, in turn, causeundesirable corrosion of steel and stainless steel, especially atelevated temperatures. Sulfuric acid, sulfurous acid and hydrogensulfide suffer serious disadvantages due to the presence of sulfur whichincreases the tendency of sulfur to exit with the final biodieselproduct. This, in turn, causes potential failure of sulfur level limitsand the formation of unwanted sulfur oxide in emissions frombiodiesel-burning engines. Arsenic acid, chromic acid, hydrocyanic acidand hydrofluoric acid are undesirably hazardous to use and/or requireunwanted additional treatment methods for the disposal of undesirablebyproducts. Lastly, iodic acid does not produce undesirableprecipitates, but it is economically not viable.

When an organic acid is used, no insoluble salt is formed and thus it isunnecessary to subject the stream to any solids separation operation.Suitable organic acid include weak organic acids, such as formic acid,acetic acid and propionic acid. In such instances, the pH of theglycerin rich stream resulting from transesterification may first beadjusted below 8.0, preferably between from about 6.5 to about 7.0.

FIG. 5 contrasts the inventive process wherein an organic acid 325 isused in the neutralization of the alkali catalyst versus a mineral acid.The use of an organic acid renders the steps of filtration, rinsing ofthe filter cake and vacuum drying unnecessary and thus offers advantagesover the use of the mineral acid.

A system may be constructed in accordance with the teachings set forthherein for the production of biodiesel from a feedstock, such as a lipidfeedstock having free fatty acids. The system may include:

(1) an optional conditioning reactor which is operative to continuouslyconvert the feedstock to a conditioned feedstock. The conditioningreactor is operative to heat, mix and filter the feedstock in order toproduce a conditioned feedstock;

(2) an optional system for continuously measuring the concentration ofthe free fatty acid in the conditioned feedstock. Suitable systemsinclude an in-line free fatty acid titration device which is operativeto quantify the concentration of the free fatty acid in the conditionedfeedstock;

(3) a glycerolysis reactor wherein the free fatty acid in the feedstockis continuously reacted with glycerin to produce a glyceride. Thisreaction may be in response to a signal from the in-line free fatty acidtitrator;

(4) a transesterification reactor for continuously reacting theglyceride with an alcohol and which is operative to convert theglyceride to a fatty acid alkyl ester and glycerin, preferably by analkali catalyzed reaction. This reaction may proceed in response to thesignal from the in-line free fatty acid;

(5) a separator for continuously separating the fatty acid alkyl esterfrom the glycerin and which is operative to produce a fatty acid alkylester rich stream and a glycerin rich stream. Suitable separatorsinclude a clarifier or a phase separation centrifuge which is operativeto produce a (first) liquid phase in which the fatty acid alkyl ester isconcentrated and a (second) liquid phase in which glycerin isconcentrated.

(6) a purifier for continuously purifying the fatty acid alkyl esterrich stream and recovering the alcohol from the fatty acid alkyl esterrich stream; the purifier being operative to produce a purifiedbiodiesel product and a first wet alcohol stream. Suitable purifiersinclude fractionation and distillation columns. In a preferredembodiment, the fatty acid alkyl ester rich stream is purified byreactive distillation;

(7) a purifier for continuously purifying the glycerin rich stream andrecovering alcohol from the glycerin rich stream; the purifier beingoperative to produce a purified glycerin product and a second wetalcohol stream. Suitable purifiers include fractionation anddistillation columns, including reactive distillation;

(8) a purifier for continuously purifying the wet alcohol streams thatis operative to produce a purified alcohol product. Suitable purifiersinclude an alcohol fractionation column for treating the alcoholstreams; and

(9) pathways for recycling at least a portion of the purified glycerinproduct to the glycerolysis reactor and recycling at least a portion ofthe purified alcohol into the transesterification reactor forcontinuously reacting with the glyceride.

Referring to FIG. 1, a preferred embodiment of a biodiesel productionprocess 10 for the conversion of high free fatty acid feedstocks intobiodiesel is presented.

In feedstock introduction step 12, the feedstock is introduced toprocess 10. The introduced feedstock is preferably conditioned infeedstock conditioning operation 14 wherein feedstock is heated andmixed in conditioning reactor 16; the high free fatty acid feedstockbeing heated and mixed to ensure a uniform mixture. The free fatty acidmay be quantified, such as in an in-line free fatty acid titration 18,wherein the concentration of free fatty acids in the feedstock ismeasured. In a first separation, solid (insoluble) substances areremoved in filter 24.

The feedstock may include at least one free fatty acid at aconcentration in the range of about 3 percent to about 97 percent byweight; moisture, impurities and unsaponifiable matter at aconcentration up to about 5 percent by weight; and a remainder thatincludes monoglycerides, diglycerides and/or triglycerides. Thefeedstock may further include trap grease.

Preferably, the conditioning step is carried out and produces aconditioned feedstock with a temperature in the range of about 35° C. toabout 250° C. and more preferably in the range of about 45° C. to about65° C. In a preferred embodiment, the feedstock is heated to atemperature in the range of about 55° C. to about 65° C. Preferably, theresulting conditioned feedstock is substantially free of insolublesolids.

The conditioned feedstock is introduced to a glycerolysis reaction at 26which preferably comprises glycerin addition step 28, heating step 32,glycerolysis step 34 in which free fatty acids are converted toglycerides and glycerolysis effluent cooling step 38.

Preferably, glycerolysis reaction step 26 further comprises performingthe glycerolysis reaction at a temperature in the range of about 150° C.to about 250° C.; and removing water from the environment of theglycerolysis reaction. More preferably, glycerolysis reaction step 26further comprises using two or more continuous stirred tank reactors inseries.

In a preferred embodiment the free fatty acid and glycerin arecontinuously reacted, typically in the absence of a catalyst, in aglycerolysis reactor at a temperature of about 220° C. and at a pressureof about 2 pounds per square inch absolute, in an esterificationreaction to produce an effluent stream that contains less than 0.5percent by weight of free fatty acids and a plurality of glycerides.Preferably, the purified glycerin product is continuously added to theglycerolysis reactor at a rate in the range of about 35 percent to about400 percent of the stoichiometric amount of free fatty acids and wateris continuously removed from the glycerolysis reactor as a vapor inwater venting step 35 through a fractionation column that returnscondensed glycerin to the glycerolysis reactor.

Preferably, the glycerolysis reactor 34 comprises at least twocontinuous stirred tank reactors that are operated in series, thereactors having a combined residence time of not greater than about 400minutes for feedstock with a 20 percent by weight free fatty acidconcentration.

Water is preferably removed as vapor through a fractionation column or adistillation column that returns condensed glycerin to the glycerolysisreactor.

The effluent from glycerolysis reaction step 26 is introduced to alkalicatalyzed transesterification reaction at 42 which preferably comprisesalcohol metering step 44, catalyst metering step 46, alkoxide additionstep 48 and transesterification step 50 wherein the glycerides undergotransesterification in the transesterification reactor.

In transesterification step 50, glycerides are contacted with aneffective amount of alcohol and an effective amount of alkali catalystunder conditions wherein the glycerides, alcohol and alkali catalystcome into substantially intimate contact. Preferably, the alkalicatalyst is selected from the group consisting of sodium hydroxide andpotassium hydroxide.

The transesterification reaction step 42 is preferably conducted at atemperature in the range of about 20° C. to about 65° C. and at anabsolute pressure in the range of about 14.5 psia. More preferably,transesterification reaction step 42 comprises conducting thetransesterification at a temperature in the range of about 25° C. toabout 65° C. and at an absolute pressure near atmospheric. In apreferred embodiment, the alcohol and alkali catalyst are mixed atprescribed rates prior to their addition to the transesterificationreaction operation.

In a preferred embodiment, transesterification reaction step 42comprises reacting the plurality of glycerides contained in theglycerolysis effluent stream with an alcohol in the transesterificationreactor. In the transesterification reactor, the plurality of glyceridesare preferably mixed with the alcohol and alkali catalyst by an agitatorand continuously reacted with the alcohol.

Preferably, the alcohol, most preferably methanol, is added to thetransesterification reactor at a rate equal to about 200 percent of thestoichiometric amount of alcohol required for the catalyzed reaction andthe alkali catalyst is added to the transesterification reactor at arate of about 0.5 percent by weight to 2.0 percent by weight ofglycerides present in the glycerolysis effluent stream. More preferably,the alkali catalyst is dissolved in the alcohol prior to theirintroduction to the transesterification reactor.

Preferably, the transesterification reactor comprises at least twocontinuous stirred tank reactors that are operated in series, saidreactors having a combined residence time of not more than about 90minutes.

The transesterification reactor effluent stream contains a plurality offatty acid alkyl esters and glycerin. The effluent fromtransesterification reaction step 42 is preferably introduced to secondseparation at 52 in which a light phase (for instance, specific gravity0.69-0.88) is separated from a heavy phase (for instance, specificgravity 0.90-1.20). In biodiesel purification step (operation) 58,excess methanol and high-boiling impurities are preferably separatedfrom fatty acid alkyl esters in the light phase and the alcohol iscollected for reuse. Preferably, separating the fatty acid alkyl estersfrom the glycerin involves using the density difference between thefirst light liquid phase and the second heavy liquid phase to separatethem.

In biodiesel purification step 56, differences in component vaporpressures are used to separate excess alcohol and high-boilingimpurities from fatty acid alkyl esters in the light phase, and thealcohol is collected for reuse.

In a preferred embodiment, second separation step 52 comprisesseparating the fatty acid alkyl esters from the glycerin in thetransesterification effluent stream in a continuous clarifier in phaseseparation step 54. Preferably, in the continuous clarifier, a firstlight liquid phase in which the plurality of fatty acid alkyl esters areconcentrated and a second heavy liquid phase in which glycerin isconcentrated are continuously separated at a temperature of about 25° C.to about 65° C. to produce a fatty acid alkyl ester rich stream and aglycerin rich stream.

Alternatively, the separation step may be a reactive distillation orfractionation column wherein the fatty acid alkyl ester and glycerin maybe separated. The transesterification effluent stream entering thereactive column contains, in addition to fatty acid alkyl esters, acertain amount of glycerin, glycerides and unreacted or non-convertiblelipid feedstock. In the reactive column, some of the glycerin reactswith unreacted fatty acids and/or fatty acid alkyl esters to formglycerides.

In preferred embodiments, the light phase is separated in fatty acidalkyl esters purification step 56. In step 56, differences in componentvapor pressures are used to separate excess alcohol and high-boilingimpurities from fatty acid alkyl esters in the first liquid phase, andthe alcohol is collected for reuse.

Preferably, purifying the fatty acid alkyl ester rich stream step 58further comprises using a distillation column to separate the fatty acidalkyl ester rich stream into a bottoms fraction, an overhead fractioncomprising primarily the alcohol, and a side stream fraction comprisinga fatty acid alkyl ester product. Preferably, the bottoms fractionproduced by the distillation column comprises impurities, unsaponifiablematerials, monoglycerides, diglycerides, triglycerides and free fattyacids. Preferably, the fatty acid alkyl ester product produced by thedistillation column meets ASTM specification D 6751. Preferably, theoverhead fraction produced by the distillation column comprisesessentially the alcohol.

Preferably, the distillation column is operated at a pressure belowabout 2 pounds per square inch absolute and at a temperature in therange of about 180° C. to about 280° C. More preferably, thedistillation column is operated at a pressure in the range of about 0.1pounds per square inch absolute to about 1 pound per square inchabsolute and at a temperature in the range of about 180° C. to about230° C. Preferably, the distillation column contains packing materialthat is operative to achieve high efficiency vacuum distillation. Morepreferably, the distillation column is packed with a structured packing.

In preferred embodiments, the heavy phase from second separation step 52is treated in catalyst separation step 62 comprising mineral acidaddition step 64, catalyst precipitation step 66 in which the alkalicatalyst is reacted with a mineral acid to produce a solid precipitate,catalyst precipitation reactor effluent filtration step 70 in which analcohol washing step 68 occurs before the alkali salt precipitate isremoved in salt recovery step 71, filtrate separation step 72 in whichthe precipitate-free filtrate is separated into two liquid phases, withthe fatty acids and fatty acid alkyl esters floating to the top and theglycerin and most of the alcohol sinking to the bottom, pHneutralization step 74 in which the pH of the glycerin is increased, andfree fatty acid recycling step 76.

Crude glycerin may be treated in glycerin purification step 80 whereinglycerin is purified by differences in component vapor pressures. Apreferred embodiment comprises distillation or fractionation step 84 inwhich the alcohol and high boiling impurities are separated from theglycerin. Glycerin decolorization step 86 comprises using a packed bedof activated carbon to remove color and odor from the distilledglycerin.

Preferably, in purifying the glycerin rich stream and recovering alcoholfrom it to produce the purified glycerin product and a wet alcoholstream, the alkali catalyst in the glycerin rich stream is reacted witha mineral acid, such as phosphoric acid or sulfuric acid, to produce aninsoluble salt having fertilizer value that is removed from the glycerinrich stream in a solids separation operation and thereafter filtered andrinsed with the alcohol.

The pH of the glycerin rich stream is adjusted to about neutral byadding a caustic alkali solution and then further purified in a glycerindistillation column that is operated at a temperature in the range ofabout 180° C. to about 230° C. and at a pressure below about 1 pound persquare inch absolute and in a decolorization column comprising a packedbed of activated carbon operated at a temperature in the range of about40° C. to about 200° C.

In a more preferred embodiment, the pH of the glycerin rich stream isadjusted to between about 6.5 and 8.0 by the addition of an acid. Anorganic acid, such as a weak organic acid, like acetic acid, propionicacid or formic acid, is then introduced to the glycerin rich stream.Salts present in the glycerin rich stream remain soluble. Thus,filtering and rinsing steps are unnecessary by use of the organic acid.

Preferably, the wet alcohol is treated in alcohol purification step 88in which water is removed from the wet alcohol. More preferably, thewater is removed by vapor pressure differences or adsorption. In apreferred embodiment, the alcohol is purified by distillation orfractionation in alcohol distillation or fractionation step 90. In apreferred embodiment, purifying the wet alcohol stream comprisesremoving water from it to produce a purified alcohol product.Preferably, the wet alcohol stream is purified in an alcoholdistillation column that is operated at a temperature in the range ofabout 60° C. to about 110° C. and at a pressure in the range of about 14pounds per square inch absolute to about 20 pounds per square inchabsolute.

In glycerin recycling step 92, glycerin is preferably recycled to step28 and in alcohol recycling step 94, alcohol is preferably recycled tostep 44. Preferably, glycerin recycling step 92 involves recycling atleast a portion of the purified glycerin product into the glycerolysisreactor for reaction with the plurality of free fatty acids in thefeedstock. Preferably, the alcohol recycling step involves recycling atleast a portion of the purified alcohol product into thetransesterification reactor for reaction with the plurality ofglycerides. The additional alcohol required for the transesterificationreaction is supplied to the alkoxide tank. Biodiesel is delivered to itsmarket in biodiesel delivery step 96 and glycerin is delivered to itsmarket in glycerin delivery step 98.

Referring to FIG. 2, a preferred embodiment of system 110 for theconversion of high free fatty acid feedstocks into biodiesel ispresented. Biodiesel production system 110 preferably comprises thesubsystems and reactors described below wherein the alcohol employed ismethanol.

In feedstock introduction subsystem 112, the feedstock is introduced tosystem 110. In a preferred embodiment, the feed material is composed ofbetween 0 and 100 percent free fatty acid content, with the remaindercomprising mono-, di- and triglycerides, moisture, impurities andunsaponifiables (MIU).

The introduced feedstock may optionally be conditioned in feedstockconditioning subsystem 114 comprising feedstock heating and mixingvessel 116 in which the high free fatty acid feedstock is heated andmixed to ensure a uniform, homogeneous mixture with uniform viscosity.The concentration of free fatty acids in the feedstock may be measuredby in-line titration device 118. The concentration is measuredcontinuously to allow continuous control of downstream process steps.

Preferably, the feed material is heated in feedstock heating and mixingvessel 116 to ensure that all of the available lipids are liquid andthat solids are suspended. Temperatures in the range of at least 35° C.but not more than 200° C. are adequate to melt the lipids, decreasetheir viscosity and allow thorough mixing of the feedstock. A jacketedstirred tank may be used to provide agitation and maintain the feedstockat increased temperature.

The conditioned feedstock may then be introduced to glycerolysisreaction subsystem 126 which comprises glycerin addition apparatus 128,input heater 132, first glycerolysis reactor 134 and second glycerolysisreactor 136 and glycerolysis effluent cooler 138. The filtered productof step 24 is combined with glycerin and subjected to conditions thatpromote the glycerolysis reaction in glycerolysis reaction subsystem126. In a preferred embodiment, these conditions include a reactiontemperature between from about 150° C. to about 250° C. and a pressurebetween about 0.1 pounds per square inch, absolute (psia) and about 30psia. A more preferred condition is a temperature of about 220° C. and apressure of about 2 psia.

Glycerin is added to the filtered grease feedstock in excess of the freefatty acid molar quantity of the grease feedstock. This excess is in therange of 10 percent to 300 percent excess glycerin (from 110 percent to400 percent of the stoichiometric amount). In this embodiment, theglycerolysis reactors used as elements 134 and 136 are configured as twoheated, continuous stirred tank reactors in series. In these vessels,the mixture of glycerin and grease (containing free fatty acids) isagitated to keep the two immiscible fluids in intimate contact.

In a preferred embodiment, mixing is provided by an agitator. Underthese conditions, the free fatty acids are converted into glycerides(mono-, di-, or triglycerides) with the production of water. The wateris vented as vapor and removed from the system together with any waterthat was initially present in the feedstock in water vapor vent 135. Thefree fatty acid content of the reactor effluent stream in this preferredembodiment of the invention can consistently be maintained at less than0.5 percent w/w.

Because of the corrosive nature of free fatty acids, the glycerolysisreactor is preferably constructed of materials resistant to organicacids.

The effluent from glycerolysis reaction subsystem 126 contains mono-,di-, and triglycerides and residual fatty acids. The glycerolysisreaction effluent is introduced to alkali catalyzed transesterificationsubsystem 142 which preferably comprises methanol metering apparatus144, potassium hydroxide metering apparatus 146, methoxide additionapparatus 148 and first transesterification reactor 150 and secondtransesterification reactor 151 in which the glycerides undergotransesterification.

In transesterification reaction subsystem 142, the glycerides aretransesterified with an alkali catalyst and a simple alcohol having 1 to5 carbons. In a preferred embodiment, the alkali catalyst is potassiumhydroxide and the alcohol is methanol. The residual free fatty acids aresaponified consuming a molar quantity of alkali catalyst about equal tothe number of moles of free fatty acid present.

The transesterification reaction is preferably catalyzed by potassiummethoxide, which is formed from the addition of potassium hydroxide tomethanol. The amount of potassium hydroxide added is preferablyequivalent to 0.5 percent to 2.0 percent w/w of the glycerides presentin the feed solution. The methanol and catalyst are combined and addedto the solution of glycerides coming from the glycerolysis reactors bymethoxide addition apparatus 148.

A 200 percent stoichiometric excess of methanol based upon the molarconcentration of glycerides is added to the reaction mixture. Uponentering each transesterification reactor 150 and 151, the two-phasesystem undergoes vigorous mixing.

Preferably, the reaction temperature is held between about 25° C. andabout 65° C. At this temperature, the miscibility of the phases islimited and mixing is required to achieve a high conversion rate. Theresidence time required is dependent on glyceride composition of thefeed (between mono-, di- and triglycerides), temperature, catalystconcentration and mass transfer rate.

Thus, agitation intensity is preferably considered in selecting aresidence time. Typically, the residence time required for greater than(>) 99 percent conversion of glycerides to alkyl esters is 20 to 30minutes.

In the transesterification reactor, the presence of potassium hydroxide,methanol, and fatty acid esters can be corrosive. In a preferredembodiment, at least two continuous stirred tank reactors in series areused. Suitable resistant materials are preferably chosen for thereactors.

The effluent from transesterification subsystem 142 may be introduced tophase separation subsystem 152 which comprise phase separation tank 154in which a light phase (for instance, specific gravity 0.69-0.88) isseparated from a heavy phase (for instance, specific gravity 0.90-1.2).The effluent streams from the phase separator are a light phase fattyacid alkyl esters comprised of methanol and alkyl esters (biodiesel), afraction of the excess alcohol and some impurities, and a heavy phase(crude glycerin) containing glycerin, alcohol, FAAEs, soaps, alkalicatalyst, a trace of water and some impurities.

Phase separation unit 154 is preferably a conventional liquid/liquidseparator, capable of separating of the heavy phase from the lightphase. Suitable phase separation units include commercially availableequipment, including continuous clarifier 154.

In biodiesel purification subsystem 156, excess methanol andhigh-boiling impurities may be separated from the fatty acid methylesters in the light phase in fractionation column 158 and methanolcollected for reuse. Preferably, purifying the fatty acid methyl esterrich stream subsystem 156 further comprises a fatty acid alkyl esterdistillation column 158 for separating the fatty acid alkyl ester richstream into a bottoms fraction, an overhead fraction comprisingprimarily methanol, and a side stream fraction comprising a fatty acidalkyl ester product.

Preferably, the bottoms fraction produced by distillation column 158comprises impurities, and unsaponifiable materials, monoglycerides,diglycerides, triglycerides and fatty acids. Preferably, the fatty acidmethyl ester product produced by distillation column 158 in FIG. 2 meetsASTM specification D 6751.

Preferably, the overhead fraction produced by distillation column 158comprises essentially methanol. Preferably, distillation column 158 isoperated under pressure below about 2 pounds per square inch absoluteand at a temperature in the range of about 180° C. to about 280° C. Morepreferably, distillation column 158 is operated under pressure in therange of about 0.1 pounds per square inch absolute to about 2 pounds persquare inch absolute and at a temperature in the range of about 180° C.to about 230° C. Preferably, distillation column 158 contains highefficiency structured packing material.

The heavy phase separated in phase separation tank 154 is preferablytreated in catalyst separation subsystem 162 comprising a mineral acid(such as phosphoric acid) addition apparatus 164, catalyst precipitationreactor 166, catalyst precipitation reactor effluent filter 170 in whichwashing with methanol 168 occurs before the potassium phosphateprecipitate 171 is removed from the filter, filtrate separation tank172, pH neutralization tank and free fatty acid recycling apparatus 176.

In catalyst separation subsystem 162, the crude glycerin phase is pumpedto a catalyst precipitation reactor where a mineral acid 164 is added.Preferably, the amount of acid added is a molar quantity equal to themolar quantity of alkali catalyst used in the transesterificationreaction. The product of the reaction is an insoluble salt that can beseparated as a solid. In addition to forming an insoluble salt, the acidconverts soaps formed in transesterification reaction subsystem 142 tofree fatty acids.

In a preferred embodiment, potassium hydroxide is used as thetransesterification catalyst, and the precipitation reaction usesphosphoric acid to form monobasic potassium phosphate. This salt is notsoluble in this system and can be removed by simple filtration. As thepotassium phosphate salt is filtered in catalyst precipitation reactoreffluent filter 170, methanol 168 is used to wash glycerin and otherprocess chemicals off of the precipitate.

The filtrate from catalyst precipitation reactor effluent filter 170 issent to another phase separation operation where two liquid phases formand separate according to their relative specific gravities in filtrateseparation tank 172. Glycerin, water, impurities and most of themethanol report to the bottom or heavy phase, while fatty acid alkylester, some alcohol and fatty acids report to the top, or light phase.The light phase is combined with the light phase from the previous phaseseparation subsystem (subsystem 152) and sent to the fractionationcolumn 158. The heavy phase is sent to a reaction operation where anyresidual acid is neutralized in pH neutralization reactor 174 by addinga small amount of caustic. In a preferred embodiment, this is performedin a continuous stirred tank reactor.

Following pH neutralization reactor 174, the crude glycerin phase issent to the glycerin refining subsystem 180, where the methanol andwater are separated and collected for further purification and theglycerin is separated from the high boiling impurities. In a preferredembodiment, glycerin separation is performed in glycerin distillation orfractionation column 184 with a glycerin side draw. The distilledglycerin may further be treated in glycerin decolorization column 186 inwhich activated carbon is used to remove color and odor from thedistilled glycerin.

The methanol recovered from the distillation column contains traceamounts of water and is therefore considered a “wet” methanol streamthat must be purified prior to reuse in the process in methanolpurification subsystem 188. This “wet” methanol stream is collected andpurified by distillation in methanol purification column 190 beforebeing pumped back into the inventory storage tanks.

The distilled glycerin stream is then subjected to decolorization anddeodorization through activated carbon bed 186. The feed enters thecolumn from the bottom and is allowed to flow upwards through theactivated carbon bed resulting in a colorless, solventless and salt freeglycerin that is >95 percent pure.

Glycerin recycling pump 192 may be used to recycle glycerin to glycerinaddition apparatus 128. Methanol recycling apparatus 194 is preferablyused to recycle methanol to methanol metering apparatus 144.

Biodiesel is then delivered to its market in biodiesel delivery vehicle196 and glycerin is delivered to its market in glycerin delivery vehicle198.

With respect to the above description then, it is to be realized thatthe optimum dimensional relationships for the parts of the invention, toinclude variations in size, materials, shape, form, function and mannerof operation, assembly and use, are deemed readily apparent and obviousto one skilled in the art, and all equivalent relationships to thoseillustrated in the drawings and described in the specification areintended to be encompassed by the present invention.

Therefore, the foregoing is considered as illustrative only of theprinciples of the invention. Further, since numerous modifications andchanges will readily occur to those skilled in the art, it is notdesired to limit the invention to the exact construction and operationshown and described, and accordingly, all suitable modifications andequivalents may be resorted to, falling within the scope of theinvention.

EXAMPLES Example No. 1

Rendered yellow grease with a free fatty acid concentration of 20percent by weight and 2 percent moisture, impurities and unsaponifiables(MIU) was fed to continuous stirred tank glycerolysis reactors at 100pounds per minute (lbs/min). The grease was filtered and titratedintermittently as it was fed to the glycerolysis reactor. Glycerin wasadded at a rate of 13 lbs/min. The temperature of the grease andglycerin mixture was raised to 210° C. as it was fed into the first ofthe glycerolysis continuous stirred tank reactors. In the reactor, thepressure was reduced to 2 psia and the temperature was maintained at210° C. The vessel was fitted with a high intensity agitator to keep theimmiscible liquids in contact. Water vapor produced by the reaction wasremoved through vents in the reactor headspace. The residence time ineach of the glycerolysis reactors was 2.5 hours. The conversion of fattyacids to glycerides in the first vessel was 85 percent. The fatty acidconcentration leaving the second reactor was maintained at 0.5 percentw/w.

The product from the glycerolysis reactors was cooled to 50° C. and fedcontinuously to the transesterification reactors in which a solution ofpotassium hydroxide in methanol was added. The potassium hydroxide wasadded at a rate of 1.1 lbs/min and mixed with 22 lbs/min of methanol.The transesterification took place in two continuous stirred tankreactors in series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters, a small amount ofunreacted glycerides and a small concentration of the unreacted methanolfloated to the top. The glycerin, the majority of the unreactedmethanol, some fatty acid methyl esters, potassium hydroxide and soapssank to the bottom.

The bottom, or heavy phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterification stepwas reacted with 1.96 lbs/min phosphoric acid. The soaps converted tofree fatty acids and the potassium hydroxide was neutralized. Theproduct of this acidification was monobasic potassium phosphate, whichwas not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out and thefiltrate was fed to a second phase separation tank where the fatty acidmethyl esters and free fatty acids present in the filtrate floated tothe top and the glycerin and methanol sank to the bottom. The top, orlight, phase was mixed with the light phase from the first phaseseparation tank and fed to the fatty acid methyl ester fractionationcolumn. The pH of the heavy phase was adjusted back to 7.5 withpotassium hydroxide and fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and18 lbs/min of glycerin. The glycerin produced was more than 95 percentpure with non-detectable concentrations of salts and methanol. Thisglycerin stream was split into two streams: 13 lbs/min was recycled backto the glycerin feed tank for the glycerolysis reaction and 5 lbs/minwas pumped through the decolorization column and collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2 lbs/min of methanol was recovered and 92lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 2

Fancy bleachable inedible tallow with a free fatty acid concentration of4 percent by weight and 0.5 percent MIU (moisture, impurities andunsaponifiables) was fed to a continuous stirred tank reactor at 100lbs/min. The grease was filtered and titrated continuously as it was fedto the glycerolysis reactors. Glycerin was added at a rate of 2.6lbs/min. The temperature of the grease and glycerin mixture was raisedto 210° C. as it was fed into the first of the glycerolysis continuousstirred tank reactors. In the reactor the pressure was reduced to 2 psiaand the temperature was maintained. The vessel was fitted with anagitator to keep the immiscible liquids in contact. Water vapor producedby the reaction was removed through vents in the reactor headspace. Theresidence time in each of the glycerolysis reactors was 2.5 hours. Theconversion of fatty acids to glycerides in the first vessel was 92percent. The fatty acid concentration leaving the second reactor wasmaintained at 0.5 percent by weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors in which a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.0 lbs/min and mixed with 22 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and a smallconcentration of the unreacted methanol floated to the top. Theglycerin, the majority of the unreacted methanol, some fatty acid methylesters, potassium hydroxide and soaps sank to the bottom.

The bottom, or heavy phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 1.79 lbs/min phosphoric acid. The soapsconverted back to free fatty acids and the potassium hydroxide wasneutralized. The product of this acidification was monobasic potassiumphosphate, which was not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out and thefiltrate was fed to a second phase separation tank where the fatty acidmethyl esters and free fatty acids floated to the top and the glycerinand methanol sank to the bottom. The top, or light, phase was mixed withthe light phase from the first phase separation tank and fed to thefatty acid methyl ester fractionation column. The pH of the heavy phasewas adjusted to 7.8 with 0.1 lbs/min potassium hydroxide and fed to theglycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and10.2 lbs/min of glycerin. The glycerin produced was more than 95 percentpure with non-detectable concentrations of salts and methanol. Theglycerin stream was split into two streams: 2.6 lbs/min was recycledback to the glycerin feed tank for the glycerolysis reaction and 7.6lbs/min was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column in which 2.1 lbs/min of methanol was recovered and93 lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) was produced.

Example No. 3

Degummed, food-grade soybean oil with a free fatty acid concentration of0.5 percent by weight and 0.5 percent MIU (moisture, impurities andunsaponifiables) was fed to a conditioning chamber at 100 lbs/min. Thegrease was filtered and titrated continuously as it was transferred fromthe feedstock conditioner. Due to the low concentration of free fattyacids, the glycerolysis section of the process was bypassed when usingthis feedstock.

The fatty acid concentration entering the transesterification reactorswas 0.5 percent by weight. The potassium hydroxide was added at a rateof 1.0 lbs/min and mixed with 22 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and a smallconcentration of the unreacted methanol floated to the top. Theglycerin, the majority of the unreacted methanol, some fatty acid methylesters, potassium hydroxide and soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 1.76 lbs/min phosphoric acid. The pH of thesolution was decreased, and the product of this acidification wasmonobasic potassium phosphate, which was not soluble in this system.

The precipitate was filtered out at 2.2 lbs/min and the filtrate was fedto a phase separation tank in which the fatty acid methyl esters andfree fatty acids floated to the top and the glycerin and methanol sankto the bottom. The top, or light, phase was mixed with the light phasefrom the first phase separation tank and fed to the fatty acid methylester fractionation column. The heavy phase was transferred to anothertank and the pH was adjusted to 7.4 with 0.1 lbs/min potassiumhydroxide. Then, the glycerin/methanol mixture was fed to the glycerinfractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and8.5 lbs/min of glycerin. The glycerin produced had a purity greater than95 percent with non-detectable concentrations of salts and methanol. Theglycerin was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.1 lbs/min of methanol was recovered and 93lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 4

Rendered trap grease with a free fatty acid concentration of 68 percentby weight and 5% MIU (moisture, impurities and unsaponifiables) was fedto the invention at 100 lbs/min. The grease was filtered and titratedcontinuously as it was fed to the glycerolysis reactors. Glycerin wasadded at a rate of 44 lbs/min. The temperature of the grease andglycerin mixture was raised to 210° C. as it was fed into the first ofthe glycerolysis continuous stirred tank reactors. In the reactor, thepressure was reduced to 2 psia and the temperature was maintained. Watervapor produced by the reaction was removed through vents in the reactorheadspace. The residence time in each of the glycerolysis reactors was3.5 hours. The conversion of fatty acids to glycerides in the firstvessel was 87 percent. The fatty acid concentration leaving the secondreactor was maintained at 0.5 percent by weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors where a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.4 lbs/min and mixed with 21 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and 10 percent of theunreacted methanol floated to the top and the glycerin, the majority ofthe unreacted methanol, some fatty acid methyl esters, potassiumhydroxide and soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 2.45 lbs/min phosphoric acid. The soapsconverted back to free fatty acids and the potassium hydroxide wasneutralized. The product of this acidification was monobasic potassiumphosphate, which was not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out at 3.1lbs/min and the filtrate was fed to a second phase separation tank wherethe fatty acid methyl esters and free fatty acids floated to the top andthe glycerin and methanol sank to the bottom. The top, or light, phasewas mixed with the light phase from the first phase separation tank andfed to the fatty acid methyl esters fractionation column. The pH of theheavy phase was adjusted back to 7.3 with 0.14 lbs/min potassiumhydroxide and fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and40 lbs/min of glycerin. The glycerin produced had a purity greater than95 percent with non-detectable concentrations of salts and methanol.This glycerin stream was recycled back to the glycerin feed tank for theglycerolysis reaction and an additional 4 lbs/min of fresh glycerin wasadded to the glycerin feed tank to provide enough glycerin feed for theglycerolysis reaction.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.1 lbs/min of methanol was recovered and 91lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 5

Rendered brown grease with a free fatty acid concentration of 37 percentby weight and 5 percent MIU (moisture, impurities and unsaponifiables)was fed to the invention at 100 lbs/min. The grease was filtered andtitrated continuously as it was fed to the glycerolysis reactors.Glycerin was added at a rate of 24 lbs/min. The temperature of thegrease and glycerin mixture was raised to 210° C. as it was fed into thefirst of the glycerolysis continuous stirred tank reactors. In thereactor, the pressure was reduced to 2 psia and the temperature wasmaintained. The vessel was fitted with an agitator to keep theimmiscible liquids in contact. Water vapor produced by the reaction wasremoved through vents in the reactor headspace. The residence time ineach of the glycerolysis reactors was 3.0 hours. The conversion of fattyacids to glycerides in the first vessel was 90 percent. The fatty acidconcentration leaving the second reactor was maintained at 0.5 percentby weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors where a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.2 lbs/min and mixed with 21 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and 10 percent of theunreacted methanol floated to the top. The glycerin, the majority of theunreacted methanol, some fatty acid methyl esters, potassium hydroxideand soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterification wasreacted with 2.13 lbs/min phosphoric acid. The soaps converted back tofree fatty acids and the potassium hydroxide was neutralized. Theproduct of this acidification was monobasic potassium phosphate, whichis not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out at 2.7lbs/min and the filtrate was fed to a second phase separation tank wherethe fatty acid methyl esters and free fatty acids floated to the top andthe glycerin and methanol sank to the bottom. The top, or light, phasewas mixed with the light phase from the first phase separation tank andfed to the fatty acid methyl ester fractionation column. The pH of theheavy phase was adjusted to 7.5 with 0.12 lbs/min potassium hydroxideand fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and25.2 lbs/min of glycerin. The glycerin produced had a purity greaterthan 95 percent with non-detectable concentrations of salts andmethanol. This glycerin stream was split into two streams: 24 lbs/minwas recycled back to the glycerin feed tank for the glycerolysisreaction, and 1.2 lbs/min was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.0 lbs/min of methanol was recovered, and89.8 lbs/min of fatty acid methyl esters meeting ASTM D 6751-02(Standard Specification for Biodiesel Fuel (B100) Blend Stock forDistillate Fuels) were produced.

Example No. 6

A feedstock containing about 0.3 weight percent of free fatty acids andabout 99.3 weight percent of glycerides (the remainder being water andinsoluble and unsaponifiable solids), at a flow rate of about 40.9pounds per hour, was heated to 50° C. and added to a solution ofpotassium hydroxide (1 percent of the feedstock flow on a weight basis)in methanol (stoichiometric ratio of 2:1 methanol:bound fatty acids inglycerides). The transesterification took place in a single continuousstirred tank reactor with a ten-hour residence time.

The transesterification effluent stream flow rate was approximately 50.3pounds per hour and consisted of approximately 79 weight percent offatty acid methyl esters, 8 weight percent glycerin, 9 weight percentmethanol, 1.6 weight percent glycerides, with the remainder being water,insoluble and unsaponifiable solids, and soaps.

This stream was separated in a flow-through separator into a light phasestream and a heavy phase stream, the light phase stream having a flow of41.5 pounds per hour and a composition of approximately 94.26 weightpercent fatty acid methyl esters, 5.6 weight percent methanol, 0.09weight percent glycerides and 0.05 weight percent free glycerin.

Free glycerin concentrations in this and the other samples in thisexample were determined using an enzyme assay solution provided bySigma-Aldrich, Inc. of St. Louis, Mo. in a kit with product code BQP-02.With this kit, free glycerin was measured by coupled, enzymaticreactions that ultimately produce a quinoneimine dye that shows anabsorbance maximum at 540 nm. The absorbance peak was measured using aBausch & Lomb Spectronic 20 spectrophotometer.

The light phase stream was analyzed for glycerin and found to containapproximately 490 ppm glycerin by weight. The light phase stream wasintroduced into a reactive distillation column maintained at 260° C. ata pressure of 150 mmHg. The overhead vapor stream from the column wascondensed, producing a liquid stream with a flow rate of about 2.1pounds per hour, consisting primarily of methanol with a glycerincontent of 135 ppm. The bottoms liquid stream, having a flow rate ofapproximately 39.3 pounds per hour, consisted of approximately 98.5weight percent fatty acid methyl esters, 1.5 weight percent glycerides,and only 3 ppm glycerin. The reactive distillation referenced in thisparagraph is schematically displayed as FIG. 6.

The gravimetric flow rates calculated using these analyses of freeglycerin in the feed to the column versus in the overhead and bottomsstreams indicated that about 98 percent of the glycerin was reacted intoother moieties in the distillation column rather than simply flowingeither to the overhead or bottoms streams.

This bottoms liquid stream was further refined to produce a biodieselstream of fatty acid methyl esters meeting ASTM D 6751-06S15 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels).

From the foregoing, it will be observed that numerous variations andmodifications may be effected without departing from the true spirit andscope of the novel concepts of the invention.

What is claimed is as follows:
 1. A process for increasing fatty acidalkyl ester yield from a feedstock containing free fatty acids, saidprocess comprising: a. reacting said feedstock in a glycerolysis reactorwith glycerin to produce glycerides and water; b. removing a vaporstream from the glycerolysis reactor; c. separating the vapor streaminto a first fraction and a second fraction, wherein the second fractionhas a high concentration of water; d. returning at least a portion ofthe first fraction to the glycerolysis reactor; and e. reacting saidglycerides in a transesterification reactor with an alcohol to producefatty acid alkyl esters.
 2. The process of claim 1 wherein the firstfraction has a high concentration of glycerin.
 3. The process of claim 1wherein the first fraction has a high concentration of feedstock.
 4. Theprocess of claim 1 wherein the first fraction has a high concentrationof feedstock and glycerin.
 5. The process of claim 1 wherein the vaporstream is removed through a vent in the glycerolysis reactor.
 6. Theprocess of claim 1 wherein the vapor stream is removed through afractionation column.
 7. The process of claim 1 wherein the vapor streamincludes at least one of entrained liquids and aerosols.
 8. The processof claim 1 wherein at least a portion of the first fraction is separatedinto a feedstock-rich stream and a glycerin-rich stream.
 9. The processof claim 1, wherein said first fraction is recovered from the vaporstream by a condenser.
 10. The process of claim 4 wherein said feedstockand glycerin are recovered together from the vapor stream to form thefirst fraction.
 11. The process of claim 9 wherein the condensercollects entrained liquids and aerosols.
 12. The process of claim 1,wherein the recovery of the first fraction occurs below the vaporpressure of water to yield the second fraction.
 13. The process of claim8, wherein said glycerin-rich stream is combined with a glycerineffluent stream from said transesterification reactor.
 14. The processof claim 13, wherein said glycerin-rich stream and said glycerineffluent stream from said transesterification reactor are furtherprocessed in a glycerin processing step.
 15. The process of claim 12,wherein the second fraction is condensed after said first fraction isseparated therefrom.
 16. A process for increasing fatty acid alkyl esteryield from a feedstock containing free fatty acids, said processcomprising: a. reacting said feedstock in a glycerolysis reactor withglycerin to produce glycerides and water; b. removing a vapor streamfrom said glycerolysis reactor, wherein the vapor stream includes water,feedstock, and glycerin; c. recovering from said vapor stream afeedstock-rich stream, a glycerin-rich stream and a water-rich stream;d. returning at least a portion of said feedstock-rich stream to saidglycerolysis reactor; and e. processing said glycerides in atransesterification process with an alcohol to produce fatty acid alkylesters.
 17. The process of claim 16 wherein at least a portion of thefeedstock and glycerin in the vapor stream is in the form of entrainedliquids and aerosols.
 18. The process of claim 16 wherein theglycerin-rich stream is further processed in a glycerin processing step.19. The process of claim 16 wherein the water-rich stream is enriched inalcohol.
 20. The process of claim 16 wherein said feedstock-rich streamand glycerin-rich stream are recovered together from the vapor stream.21. The process of claim 16 further comprising a phase separation unitfor separating at least a portion of the feedstock-rich stream from atleast a portion of the glycerin-rich stream before returning the atleast a portion of the feedstock-rich stream to the glycerolysisreactor.
 22. A process for increasing fatty acid alkyl ester yield froma feedstock containing free fatty acids, said process comprising: a.reacting said feedstock in a glycerolysis reactor with glycerin toproduce glycerides and water; b. removing a first vapor stream from saidglycerolysis reactor, wherein the first vapor stream includes water,feedstock, and glycerin; c. recovering a portion of the feedstock andglycerin from the first vapor stream creating a second vapor streamwhich includes water, feedstock, and glycerin; d. returning at least aportion of the feedstock and glycerin recovered from the first vaporstream to the glycerolysis reactor; e. recovering feedstock and glycerinfrom the second vapor stream creating a third vapor stream enriched inwater; f. separating the feedstock and glycerin recovered from thesecond vapor stream in a separation unit into a glycerin-rich phase anda feedstock-rich phase; g. returning at least a portion of thefeedstock-rich phase to the glycerolysis reactor; h. processing saidglycerides in a transesterification process with an alcohol to producefatty acid alkyl esters.
 23. The process of claim 22 wherein at least aportion of the feedstock and glycerin in the first vapor stream is inthe form of entrained liquids and aerosols.
 24. The process of claim 22wherein the glycerin-rich phase is sent to a glycerin processing unitfor further processing.
 25. The process of claim 22 wherein the thirdvapor stream is condensed to yield a water-rich liquid stream.
 26. Theprocess of claim 22 wherein the third vapor stream includes alcohol. 27.A process for increasing the glyceride content of a feedstock containingfree fatty acids, said process comprising: a. reacting said feedstock ina glycerolysis reactor with glycerin to produce glycerides and water; b.removing a vapor stream from said glycerolysis reactor, wherein thevapor stream includes feedstock, glycerin, and water; c. recoveringfeedstock and glycerin from said vapor stream; and d. returning at leasta portion of said recovered feedstock to said glycerolysis reactor. 28.The process of claim 27 wherein at least a portion of the recoveredglycerin is further processed in a glycerin processing step.
 29. Theprocess of claim 27 wherein said feedstock and glycerin are recoveredtogether from the vapor stream to create an intermediate stream.
 30. Theprocess of claim 29 wherein at least a portion of the intermediatestream is returned to the glycerolysis reactor.
 31. The process of claim29 wherein at least a portion of the feedstock is recovered from theintermediate stream.
 32. The process of claim 27 wherein said glyceridesare processed in a transesterification process with an alcohol toproduce fatty acid alkyl esters.